planta electro

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The INL is a U.S. Department of Energy National Laboratoryoperated by Battelle Energy Alliance

INL/EXT-05-005

Alternate Anode

Reaction for Copper

Electrowinning

Gerald L. May

Bill Imrie

Sharon Young

Brent Hisky

Jan Miller

Michael Free

July 2005


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INL/EXT-05-00539

Alternate Anode Reaction for Copper Electrowinning

Gerald L. Maya

Bill Imrieb

Sharon Youngc

Brent Hiskyd

Jan Millere

Michael FreeeaBEA

bBechtel

cVersitec

dUniversity of Arizona

eUniversity of Utah

July 2005

Idaho National LaboratoryIdaho Falls, Idaho 83415

Prepared for theU.S. Department of EnergyOffice of Nuclear Energy

Under DOE Idaho Operations OfficeContract DE-AC07-05ID14517


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SUMMARY

This report describes a project funded by the Department of Energy, withadditional funding from Bechtel National, to develop a copper electrowinningprocess with lower costs and lower emissions than the current process. This newprocess also includes more energy efficient production by using catalytic-surfaced anodes and a different electrochemical couple in the electrolyte,providing an alternative oxidation reaction that requires up to 50% less energythan is currently required to electrowin the same quantity of copper.

This alternative anode reaction, which oxidizes ferric ions to ferrous, withsubsequent reduction back to ferric using sulfur dioxide, was demonstrated to betechnically and operationally feasible. However, pure sulfur dioxide wasdetermined to be prohibitively expensive and use of a sulfur burner, producing12% SO2, was deemed a viable alternative. This alternate, sulfer-burning processrequires a sulfur burner, waste heat boiler, quench tower, and reaction towers.The electrolyte containing absorbed SO2 passes through activated carbon toregenerate the ferrous ion. Because this reaction produces sulfuric acid, excessacid removal by ion exchange is necessary and produces a low concentration acidsuitable for leaching oxide copper minerals. If sulfide minerals are to be leachedor the acid unneeded on site, hydrogen was demonstrated to be a potential

reductant.

Preliminary economics indicate that the process would only be viable ifsignificant credits could be realized for electrical power produced by the sulfurburner and for acid if used for leaching of oxidized copper minerals on site.


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CONTENTS

SUMMARY.................................................................................................................................................iii

1. INTRODUCTION.............................................................................................................................. 1

2. PROJECT DESCRIPTION ................................................................................................................1

3. REDUCTANTS.................................................................................................................................. 3

3.1 Sulfur Dioxide .......................................................................................................................3

3.2 Sulfur.....................................................................................................................................3

3.3 Hydrogen...............................................................................................................................3

4. FERROUS REGENERATION .......................................................................................................... 5

5. ACID REMOVAL AND RECOVERY ............................................................................................. 8

6. ANODES/CELL DESIGN................................................................................................................. 9

7. PILOT PLANT OPERATION ......................................................................................................... 10

8. DESIGN CRITERIA........................................................................................................................ 10

9. ECONOMICS................................................................................................................................... 10

10. CONCLUSIONS .............................................................................................................................. 10

11. SOURCES OF COST SHARE AND FINANCING ........................................................................ 12

FIGURES

1. Simplified diagram for economic model. ........................................................................................... 2

2. Process Flowsheet. ............................................................................................................................. 7

3. Material balance for the larger plant................................................................................................. 11


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Alternative Anode Reaction for Copper Electrowinning

1. INTRODUCTION

Solvent extraction and electrowinning (SX/EW) are among the most important processes used inthe copper industry and are key components in producing some of the best quality copper in the market.

The conventional copper electrowinning process uses the water hydrolysis reaction as the anodicsource of electrons. However, this reaction generates acid and oxygen gas. The oxygen, which leaves thesystem, is associated with some acid from the electrolyte resulting in an acid mist. This mist posespotential health hazards near the electrowinning process. The acid generated also presents corrosionproblems. The cell voltages applied for this process are high and because energy consumption is directlyproportional to cell voltage, the energy consumption is also high.

The goal of the project documented in this report was to develop a process with lower costs andlower emissions that would include more energy efficient production by using catalytic surfaced anodes

and a different ionic couple in the electrolyte, providing an alternative oxidation reaction that requires upto 50% less energy to electrowin the same quantity of copper. An additional goal was to have the abilityto readily retrofit this process into existing commercial copper electrowinning tankhouses as well asincorporate it into new installations

The project resulted in a proposed process that will function without producing acid mist, thussignificantly reducing corrosion of buildings and equipment, and requiring less maintenance. The newprocess will also improve the safety and work environment, and eliminate worker exposure, as it producesno oxygen, acid mist, or other gaseous emissions. The project included a demonstration of the newprocess at large pilot-plant scale. However, preliminary economics indicate that the process would onlybe viable if significant credits could be realized for electrical power produced by the sulfur burner and foracid if used for leaching of oxidized copper minerals.

This report includes descriptions and analyses of the project, including the proposed process andthe process economics as well as conclusions drawn from the project data.

2. PROJECT DESCRIPTION

The proposed technology will employ a reaction that relies on the ferrous/ferric/sulfur dioxide(FFS) chemical species in the copper electrowinning process. The original project planning calls for theFFS technology to be developed for commercialization, focusing on production of sulfur dioxide forferric reduction, reaction of the resulting gas with the ferric ions to produce acid, removal and recovery ofthe excess sulfuric acid, and development of anode materials and electrolyte flow manifold design. Theresearch project included a large pilot-plant scale process demonstration at a commercial electrowinningplant. The energy required to produce copper using the FFS process instead of present electrowinning

technology was quantified as part of the project.The FFS process involves the addition of ferrous sulfate to conventional copper electrolyte and the

use of catalytic anode surfaces. The anode reaction during electrowinning then becomes the Fe2+/Fe3+

couple. Separate from the electrowinning cell, ferrous ions are regenerated using gaseous SO2 andforming H2SO4 as a potentially valuable by-product (e.g., to use onsite for leaching of ore). The diagramin Figure 1 outlines the FFS process.


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LeachArea 200

Solvent Extraction

Area 300

Electrowinning

Area 500

SO2

absorption

Tank Farm

Area 400

Lean

Electrolyte

Cathode Copper

Lean Electrolyte

Mine

Elemental Sulfur

Activated

Carbon

Modules

Electric Power

Sulfur Burner,

Boiler & Generator

APU

Modules

Raffinate

Acid

Concentration,

Tanks and

Loading

Surplus Acid

Ferrous RegenerationArea 510

Acid Recovery

Area 520

Cogen Power

Heat

Nearby Oxide

Leach

Off-gas

Cleaning

ALTERNATIVE ANODE REACTION FOR COPPER ELECTROWINNING

Figure 1. Simplified diagram for economic model.

The excess acid recovered is quite dilute compared to commercial grade acid and must be used inthe immediate vicinity of the electrowinning plant for leaching of additional copper. Excess acid is onlyconsumed with the leaching of oxidized copper minerals in alkaline host rock. Host rock in the

southwestern USA is usually quite alkaline and can readily consume the excess acid produced. Howeverwhen sulfide copper minerals are leached, the leaching process itself produces acid. The cost of producingand recovering unusable acid is therefore a fatal flaw for the economics of the process. The projecttherefore included investigating other potential reductants as well as the potential of upgrading thesulfuric acid to a salable product. However, the project was based on the FFS process.


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3. REDUCTANTS

3.1 Sulfur Dioxide

Sulfur dioxide is a versatile compound; in the presence of a strong oxidant it plays the role of a

reducing agent, while in the presence of a strong reductant it acts as an oxidizing agent. Since ferric ionsare stronger oxidants than sulfur dioxide, this oxide of sulfur is good for the reduction of ferric ions in thesolution and is the reaction that provided the basis for the original process development. The reactionproceeds quickly in the presence of activated carbon as a catalyst, most likely in much less than twominutes. The overall reduction reaction using sulfur dioxide is:

2Fe3+ + SO2 + 2H2O = H2SO4 + 2H+ + 2Fe2+

The above reaction also indicates that one mole of SO2 is needed to remove one mole of copper.One mole of excess H2SO4 is generated per mole of copper removed from the solution.

The effect of oxygen on this reduction mechanism was evaluated due to the presence of significantamounts of residual oxygen in SO2 gas streams produced by sulfur burners. The gas mixtures used inthese tests were composed of sulfur dioxide and oxygen (total 20% by volume) with the balanceconsisting of nitrogen.

Results of sulfur dioxide and oxygen tests showed that some the ferric ions are regenerated if thegas contains oxygen. The predicted results showed that 21% of the sulfur dioxide consumed will beutilized to counteract the effect of oxygen when the gas mixture consistes of 12.37% SO2 + 7.96% O2with the balance consisting of nitrogen. This value reduces to 7.69% with a gas mixture of composition17.91% SO2 + 2.01% O2 with the balance consisting of nitrogen.

3.2 Sulfur

The overall reduction reaction using elemental sulfur can be expressed as:

6Fe3+ + So + 4H2O = H2SO4 + 6H+ + 6Fe2+

Elemental sulfur can give up to six electrons, which is sufficient to reduce six ferric ions. Since twoFe3+ are required to remove one Cu2+, each elemental sulfur atom will be sufficient to enable the

production of three copper atoms. This reaction also shows that the excess H+ produced is one mole ofacid for each three moles of Cu produced or one-third of the excess acid produced by the SO2 reduction.

During the study of sulfur as a reducing agent, orthorhombic and liquid sulfurs were used. Theresults show that sulfur reduces the ferric ion concentration from 1.5 g/L to approximately 0.04 g/L in twohours at 100C, or to 0.1 g/L in four hours at 50C. Thus, sulfur offers significant potential as a ferric ionreductant.

3.3 Hydrogen

Hydrogen is used for reduction of metal ions in many metallic systems. It is a logical choice as acandidate for ferric ion reduction. However, there is an explosion risk associated with the hydrogen gas.The overall ferric ion reduction reaction using hydrogen gas is:

2Fe3+ + H2 = 2H+ + 2Fe2+


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The application of hydrogen as a reductant is widely known. In this study, the reduction of ferricions with hydrogen was investigated using various catalysts such as graphite and platinum. Specifically,various forms of graphite (natural, synthetic, and exfoliated) and platinum (amorphous and crystalline)particles were evaluated at different concentrations and size distributions.

One of the most effective surfaces evaluated in this study for the reduction of ferric ions in the

presence of hydrogen gas is crushed graphite electrode particles. The normalized reduction rate for theseparticles in the 20 60 mesh sample that was present in electrolyte solution containing 5 g particles/Lconcentration is 1.49 grams of ferric ions reduced per liter of electrolyte per hour per square meter ofparticle surface per liter of hydrogen gas used.

Natural graphite flakes and exfoliated graphite particles were also used as potential catalyticsurfaces to enhance the ferric reduction process with hydrogen gas. However, all testing with thesesubstances was subject to inadequate mixing due to surface hydrophobicity that caused the material tofloat when exposed to the injected argon-hydrogen gas mixture. Consequently, effective contact betweenthe electrolyte and the surfaces was minimal and the reduction rates were also negligible. No surfactantwas added to enhance the wettability for these tests.

Colloidal graphite was evaluated with and without surfactant present to aid in particle wetting.However, because the colloidal graphite is extremely hydrophobic, the introduction of 2% hydrogen inargon gas causes flotation to occur and the colloidal graphite does not have sufficient wettability tosustain a significant ferric ion reduction rate.

Synthetic graphite in various sizes was evaluated as a potential catalytic surface for ferric ionreduction. Graphite particles (-200 mesh) produced synthetically from desulfurized petroleum coke wereevaluated and found to produce significant ferric ion reduction in the copper electrowinning electrolyte.The normalized rate of reduction was approximately 0.5 grams of ferric reduced per liter of solution perhour per square meter of catalytic surface area per liter of hydrogen gas.

Graphite particles (-20+100 mesh) were also tested. The larger particles also gave significant

reduction rates of between 0.059 and 0.8 grams of ferric reduced per liter of solution per hour per metersquared of external surface area per liter of hydrogen gas.

Thus, hydrogen gas was found to carry out the ferric reduction reaction effectively even when usedin the very small concentration of 2% by volume. The presence of graphite as a catalyst for this reactiongreatly improves the rate. Extrapolation from bench scale tests suggested that hydrogen reduction of ferricions might be feasible on a commercial scale. Graphite and platinum were found to be the most effectivecatalysts in carrying the ferric ion reduction by hydrogen gas. Platinum catalyzed the process when usedin amorphous form as well as in the form of a supported catalyst on graphite.

Platinum-coated graphite was evaluated to determine if the presence of platinum on the graphiteparticles enhanced the ferric ion reduction process in the presence of 2% hydrogen gas in argon. Theplatinum was coated by electrolysis deposition. However, a complete coating was not achieved. The testresults showed a significant normalized rate of reduction occurred (2.58 grams of Fe3+ reduced per liter ofsolution per hour per meter square of catalyst surface area per liter of hydrogen gas).

The cost analysis for the use of ferric/ferrous reaction along with hydrogen as a reductant isdiscussed below. The electrical energy saved by shifting from water hydrolysis to ferrous/ferric reactioncan be calculated as


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.

...26800

wM

nVE

where E is electrical energy saved in kWh/ton of copper, n is number of electrons, Mw is the atomicweight of copper, V is the voltage reduction and is the current efficiency.

For a voltage reduction of 1.0V operating at a current efficiency of 95%, the energy savings willbe:

lbCukWhrtonCukWhrxx

E /332.0/98.66595.054.63

20.126800

In the United States, on a national level, the average price for electrical energy in 2001 was 7.32cents/kWhr1 and the average price for electricity for the industrial sector was 5.04 cents/kWhr2. Hence,using this data, the total savings obtained by shifting from water hydrolysis to ferrous/ferric reaction is1.67 cents/lb of Cu.

The hydrogen required for the process can be obtained from an onsite natural gas reformer, whichwill cost around 34 cents/lb of H2

3. Since one mole of hydrogen will give one mole of copper at cathodein electrowinning, the investment for the onsite supply of hydrogen will be 1.08 cents/lb of Cu.

The difference between the cost savings using hydrogen (1.67/lb) and the cost of hydrogen(1.08/lb) is 0.59/lb of Cu.

4. FERROUS REGENERATION

Sulfur dioxide has been shown to react with ferric ion in copper electrolyte according to thereaction:

2H2O + SO2 + 2Fe3+

2Fe2+

+ SO42-

+ 4H+

The electrolyte typically contains 160 g/L H2SO4, 40 g/L Cu2+, 27 g/L Fe2+, and 3 g/L Fe3+. The

reaction is catalyzed by activated carbon.

When combined with copper electrowinning using ferrous oxidation as the anode reaction, theoverall reaction becomes:

2 0 3+ 2-

4 4

0

4 2 2 2 4

Electrowinning: CuSO + 2Fe + Cu + 2Fe + SO

Overall Reaction: CuSO + 2H O + SO Cu + 2H SO

A key to making this process work is reacting SO2 with Fe3+ in copper electrolyte. Bench and pilot-

scale work passed a pure SO2 gas/electrolyte mixture through a module of activated carbon, where aquick conversion of SO2 to H2SO4 was achieved. The use of pure SO2 to reduce the ferric ion has beendemonstrated to be uneconomical. On an industrial scale, producing SO2 from sulfur in such a way thatthe SO2 can be utilized in this process is the issue. The SO2 can be produced from burning sulfur in air.The SO2 must then be transferred from the resulting nitrogen-oxygen-SO2 gas mixture into copperelectrolyte, and passed through the activated carbon. As discussed above, sufficient oxygen will beabsorbed into the electrolyte to oxidize Fe2+, making the process less efficient by about 20%. A provision


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to scrub the large volume of nitrogen/oxygen gas mixture to remove residual SO2 before emitting to theatmosphere must be considered.

The goal for a 175,000 ton/year copper production plant is to emit less than 40 tons SO2/year intothe atmosphere. Since approximately one ton of SO2 is required for each ton of copper produced, less than0.02% of the SO2 produced can be emitted or less than an overall average of about 25 ppm in the exiting

gas if the burner gas contains 12% SO2.

The portion of the flowsheet to be developed consists of the introduction of molten sulfur to thesulfur burner, producing gas containing SO2 to be introduced into the electrolyte and reaction of the SO2with the ferric ion to produce ferrous ion and acid. The flowsheet is shown in Figure 2.

A sulfur burner can provide a wide range of SO2 strengths, typically from 8-12%. A higher strengthgas has the advantage of less air required thus increasing the available heat, reducing excess oxygen, andreducing the amount of inert gas to be treated prior to exiting the system. Less oxygen reduces thepossibility of oxidizing a portion of the ferrous ion outside of the ferric reduction reaction. However, ifthe concentration of the SO2 is much above 12%, the gas becomes very corrosive, in part due to theformation of some sulfur trioxide. The formation of the trioxide also has a significant negative effect on

the economics as more sulfur is required to provide sufficient sulfur dioxide and more acid must beremoved from the electrolyte. Different and much more expensive materials of construction are alsorequired for the plant. Therefore the flowsheet was developed around a gas concentration of 12% SO2.

The hot gas exiting the sulfur burner will pass through a waste heat boiler to produce steam forpower generation and secondary steam suitable for process uses. The gas will then go to a humidifying orquench tower that will further cool the gas and bring it to near saturation with water to avoidcrystallization of the electrolyte.

The SO2 bearing gas and the high ferric lean electrolyte will then enter reaction towers where theSO2 is absorbed into the electrolyte. The gas exiting the towers will pass through a demister to ensure thatenvironmental regulations are met. The reaction towers may be conventional towers, 10 each of 10 ft.

diameter, each equipped with 12 feet of packing, operated in parallel for both gas and electrolyte flows.The budgetary price for the reaction towers alone is $1,700,000 25%.

A potential alternative is Monsantos DynaWave reactors in which the gas is treated in two countercurrent reactors before being demisted. Without a contract assuring confidentiality and agreement topurchase their product, Monsanto would not provide costs of such reactors but indicated that the sulfurburner, waste heat boiler, humidifying tower and reaction towers should be order-of-magnitude$7,000,000 installed.

The electrolyte will then pass through activated carbon modules. The design was based on a five-minute retention time. Several vendors were contacted and they indicated they could not supply aneffective system (with estimates of from 33 to over 150 vessels being deemed necessary) with enoughcapacity for the 5-minute retention time for 66,000 gpm. One vendor suggested he could provide a non-conventional system, but would require an agreement that he would provide the commercial unit(s).Nichem was the only one to indicate they thought what was requested was feasible. They worked withanother company, AFT, to come up with a design and budgetary prices for the activated carbon modules,resulting in two 80 ft diameter mix tanks, constantly stirred by paddles, and each containingapproximately 60,000 lb. of activated carbon. Spent carbon is removed from these tanks by vibratingscreens. The spent carbon is reactivated by a rotary kiln and is fed to a feed tank for make-up to the mixtanks. To minimize attrition, it is recommended that a coconut shell-based carbon be used instead of a


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gastoatmos

phere

gas

electrolyte

66,000gpm,160g/lacid,

38g/lCu

2+,27.6g/lFe2

+,2.4g/l

Fe3

+

Sulfur

gas

SO2@~1

2%,~8%O2

electrolyte

670lbSO2/minute

0.15gplfe

rric

electrolyte

164g/laci

d

w

ater

washwater

toraffinate

Reaction

Tower(s)

2-stageCountercurrent

Scrubber

Activated

carbon

modules

Carbon

Reactivation

Demister

Sulfur

Burner

Waste

heat

boiler

Humidifying

tow

er

Figure2.ProcessFlowsh

eet.


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coal-based carbon. Following carbon contact, the electrolyte will need to be filtered to remove carbonfines with DE precoat rotary vacuum filters. Ten (10) filters are required for the 66,000-gpm-electrolyteflow. The electrolyte equipment is estimated at $3,900,000 25% and includes the following:

One initial activated carbon fill (120,000 lb)

Two ea. 80 ft diameter mix tanks, site built concrete

Two ea. 200 HP low speed bridge mount mixers

Two ea. 60 ft. diameter strip/recycle tanks

One 15 ft diameter spent carbon tank

One direct fired rotary kiln

One 8 ft diameter quench tank with fines carryover

One lot process water recycle system from contactors

One lot 1,200 ft 24 in. diameter carbon steel piping and flanges

1 lot process water valves of varying diameters

10 ea. stainless steel rotary vacuum filters

The retention time in the carbon modules can probably be reduced to less than two minutes fromthe five minutes originally used for sizing. Decreasing the capital costs of the carbon modules and carbonregeneration system by 50% gives an equipment cost of $2M. A factor of 2.5 is used to convertequipment costs into installed costs of $5M.

The capital costs for the ferrous regeneration system are given below. The annual carbonreplacement costs are estimated to be 120,000 lb x 20% x $0.60 or about $15,000.

Installed

Sulfur burner, WHB, Quench Tower,

Reaction Towers and Demister $7M

Activated Carbon Modules and Filters $5M

Total $12M

5. ACID REMOVAL AND RECOVERY

It was necessary to evaluate, assess, and recommend processes and equipment required toeffectively remove sulfuric acid produced from the sulfur dioxide reduction step of the process andprepare the installed costs of said equipment. Secondly, it was necessary to provide an effective pilot unitfor plant acid retardation test work. Third, equipment was identified that could be used to upgrade theextracted sulfuric acid to a concentration of greater than 300 g/L.


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Three suppliers were identified that can provide equipment for the removal of sulfuric acid fromthe process stream. The recovery and cost of the three vary significantly. The acid recovery ranges from72.5% to 94.2% of the input acid and installed costs range between approximately $29 and $42 milliondollars. The table below summarizes the acid recovery as well as the copper and iron losses for each pieceof equipment.

Acid Recovery Copper Loss Iron Loss

Higgins Loop 94.2% 4.2% 6.4%

MPT/Ari 85.0% 12.8% 12.9%

EcoTec 72.5% 11.1% 12.0%

Each of the three suppliers have pilot plants that are available for onsite rental testing. The reasonfor determining the efficiency and cost of the full sized plant was to identify one system, which was morecost effective than the others and then obtaining a pilot plant unit that would verify the stated results.Rental costs ranged between $50,000 and $250,000. During the course of identifying pilot plants, Phelps

Dodge purchased individual components for use in their test facility and proceeded forward with theirown pilot test work from which no test results are available.

The discharge acid concentration from the three pieces of equipment, previously stated, rangedbetween 110 and 193 g/L, far below the requested 300 g/l. Various evaporators and membrane processeswere investigated to determine if the concentration to 300 g/l was possible to obtain.

Nine different evaporator types and manufacturers were contacted and all but two said that theunits for which information was requested were larger than they had ever made. All who had experiencewith submerged combustion burners stated that this process is not the correct application for thesubmerged flame and there would be additional costs associated with environmental containment of thedischarge. The most practical unit is a two-stage falling film MVR (mechanical vapor recompression)evaporator, which will produce the required acid concentration and provide purified water back to theprocess.

Three different membrane units were identified but none were capable of producing the 300 g/lacid concentration required. The problem is that the viscosity of the acid solution changes and wouldbecome more like honey or thick syrup and would not pass through the membrane. Heat would berequired to allow the solution to pass through the membrane and Hastalloy C or Carpenter 20 would bethe material of construction making the units too costly for commercial installation.

It was determined that the membrane equipment could be used cost effectively to recover copperand iron that would pass through the acid removal unit and out of the process flow unless a secondarymetal recovery step was installed. Estimates indicated that 98% of the copper and 97% of the iron couldbe recovered from the discharge stream.

6. ANODES/CELL DESIGN

Earlier testwork and experience of some of the team members indicated that anode and cell designsneeded significant improvement. At the outset of the project, Phelps Dodge indicated that they would dothis development work and considered it proprietary and confidential. Phelps Dodge personnel indicatedthat they had been successful. No other information was provided.


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7. PILOT PLANT OPERATION

Phelps Dodge Morenci has a four-cell pilot plant, which was to be used as the demonstration plant.Their personnel indicated that they considered its operation and the information derived to be proprietaryand confidential. No data or information was provided other than that they had had a successful pilot plantrun of several months and achieved the expected decrease in cell voltage. No comments were made

concerning the cell design nor offset costs of electrolyte pumping.

8. DESIGN CRITERIA

A flowsheet with material balance was developed for a plant producing 175,000 TPY and for asmaller plant producing 50,000 TPY. A rough draft of cell design criteria was prepared for both.Although the project originally was concerned with the larger plant, it is much more likely that a firstcommercial operation would be the smaller plant size. The material balance for the larger plant ispresented in Figure 3.

9. ECONOMICS

Preliminary calculations of Net Present Value relative to conventional electrowinning gave anegative value of about $45M. No credit was given to operating costs for electrical power generated bythe sulfur burner or for acid that can be used for leaching. Both of these items are highly dependent on thelocation and type of ore body where the EW plant would be located.

10. CONCLUSIONS

No technical or operational fatal flaws were discovered.

Sulfur dioxide regeneration of ferrous ions would be the most likely first commercialization iflocated at an oxide leach property where power costs are high and sulfuric acid availability islimited, giving the best economics. This process is not viable for leach systems that do not requireda significant amount of sulfuric acid, such as sulfide mineral leaching.

Hydrogen is an interesting alternative reductant that should be investigated more thoroughly,especially for sulfide leach applications. This reductant produces no excess acid and wouldtherefore be useful in leaching of all types of copper minerals.

Sulfur can also act as a reductant. It also produces excess acid but only about one third as much asthe sulfur dioxide reductant.

Using sulfur dioxide produced in a sulfur burner at 12% SO2 requires a sulfur burner, waste heatboiler, quench tower, reaction towers, and demister to introduce the reductant into the electrolyteand vent the waste gases to the atmosphere meeting environmental regulations. The residualoxygen will reoxidize ferrous ions causing as much as a 20% loss in efficiency.

The electrolyte containing the absorbed sulfur dioxide is passed through activated carbon beds

nominally for five minutes. The retention time required is expected to be much less than twominutes.

The excess acid produced can be recovered by ion exchange and can produce acid ranging from110 to 190 g/L depending on the supplier and configuration of equipment.


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Hi

Ferr

ic

Flows

hee

t

62,319gpm

Ferrous

Regenera

tion

Line

Ac

idRecovery

Line

176,248tpy

62,319gpm

Cu

38g/L

Irontotal

30

61,226gpm

Ferric

0.15

1458

gpm

1458gp

m

Ferrous

29.85

269,880

tpyH2SO4

Acid

162

Acid50g/L

1094gpm

1094gpm

Rich

Elec

tro

lytefrom

SX

Cu

39.3g/L

Iron

total

30

Ferric

0.15

Ferrous

29.85

Acid

158

Electro

winningCells

Lean

Elec

tro

lyte

Cu

38g/L

Irontotal

30

Ferric

2.4

Ferrous

27.6

Acid

160

SO2Source

Activated

Carbon

Modules

APUModules

Ra

ffina

te/wa

ter

Cu

0.2g/L

Irontotal

2.8

Ferric

2.8

Ferrous

0

Acid

5

Lea

nElec

tro

lyteTan

ks

toS

X

Cu

38g/L

Iron

total

30

Ferric

0.15

Ferrous

29.85

Aci

d

160

Ac

idStorage

Tan

ks

Cu

4.2g/L

Irontotal

3.6

Ferric

Ferrous

Acid

122g/L

C

opper

1

75,000tpy

Figure3.Materialbalanc

eforthelargerplant.


18/18

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The most likely method of upgrading acid to 300 g/L is by two-stage falling film mechanical vaporrecompression evaporation.

Membrane units can be used cost effectively to recover copper and iron that would be lost with theacid.

Preliminary economic evaluations indicated that the process would not produce a positive net

present value unless significant operating cost credits could be realized by power generation at thesulfur burner and by offsets to leaching acid costs where acid is in short supply and/or veryexpensive.

11. SOURCES OF COST SHARE AND FINANCING

Project Partner In-Kind Contributions Cash Contributions

Bateman Engineering $95,000

Bechtel International $10,000 $120,000

Phelps Dodge Mining and Phelps Dodge Miami >$3,000,000a

Department of Energy $366,963

Total >$3,105,000 $486,963

a. Phelps Dodge would not provide exact expenses.

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